WO 2009/086908 PCT/EP2008/010990 1 Process and device for generating middle distillate from 2 hydrocarbonaceous energy sources 3 4 The invention relates to a process and a device for 5 generating middle distillate from hydrocarbonaceous energy 6 sources. 7 8 It is known from the prior art to release the fuels in the 9 form of hydrocarbons contained in residues, not by reaction 10 with oxygen via combustion or gasification, but to release 11 them in material form by catalytic treatment in the absence 12 of air in an oil bath and to obtain them as a valuable 13 material. This serves for avoiding the formation of CO 2 in 14 disposal of residues and serves for producing fuels or 15 propellants from the residues. 16 17 The residue in the form of renewable raw materials, such as 18 wood and plant parts, of waste products, such as plastics, of 19 animal and plant wastes, of waste oils and other organic raw 20 materials which contain a preferably high proportion of 21 hydrocarbons and, because of their energetic utilizability 22 can be termed materials of value or energy sources, remains 23 in the oil bath until, by molecular dehydration, molecular 24 polymerization and molecular shortening (depolymeriz 25 ation/conversion into oil) these hydrocarbons can be 26 separated off as hydrocarbon vapor. 27 28 DE 100 49 377 C2 discloses a process for converting plastics, 29 fats, oils and other hydrocarbonaceous wastes into oils, 30 wherein a catalyst of sodium aluminum silicates is mixed in a 31 circulation evaporator in a circuit in a high-boiling 32 hydrocarbon, such as thermal oil, base oil or bunker-C oil, 33 and in the reactor part below the distillation system, 34 plastics, fats, oils and other hydrocarbonaceous wastes are 35 added. The reaction site for the conversion reaction into oil 36 is a circulation evaporator system which consists of a tube 37 bundle evaporator which is heated with flue gas and a reactor WO 2009/086908 PCT/EP2008/010990 2 1 connected to two tubes, which reactor needs feed and 2 discharge functions. On the reactor is arranged a 3 distillation column which takes up the catalytically cracked 4 product in vapor form and separates it into the actual 5 product diesel, a fraction for petrol production and reflux 6 into the reactor for a further catalytic cracking reaction. 7 By combustion below the circulation evaporator, hot flue gas 8 is generated and passed through the flue gas tubes of the 9 circulation evaporator. In the circulation evaporator the hot 10 flue gases cool down, wherein in the lower part of the 11 circulation evaporator temperatures of approximately 430 to 12 470 result on the inside of the tubes, where the catalyst 13 containing oils pass onto the tubes together with the molten 14 residues, which leads to a selective catalytic cracking of 15 the residues to form a hydrocarbon vapor. 16 17 The high temperature of the hot flue gases leads to the 18 formation of reaction coke which reacts with the sodium-doped 19 aluminum silicate to form a nonreactive residue which fouls 20 the plant and brings the reaction to a stop. This reaction 21 mixture of the catalyst and the reaction coke combines with 22 the walls of the circulation evaporator and the reactor to 23 form a hard residue and requires a high expenditure on 24 cleaning in short maintenance intervals. Economic operation 25 of the known process is therefore only possible with 26 restrictions. Furthermore, only low yields are achieved of 27 the heating value of the input substances. 28 29 EP 1 538 191 Al discloses a process for generating diesel oil 30 from hydrocarbonaceous residues in an oil circuit with solids 31 deposition and product distillation for the diesel product, 32 wherein the main energy application and thereby the main 33 heating proceeds via one or more pumps and wherein the flow 34 energy of the pump is braked by a counterrotating stirrer and 35 is intended to be converted into heat. Active energy input 36 via heating through the wall is not provided in this process. 37 Instead, the heat is not transported through the wall, but is WO 2009/086908 PCT/EP2008/010990 3 1 liberated directly in the reaction system. The stirrer in 2 this case serves also for complete cleaning of the surfaces 3 arranged in the circuit. Industrial conversion of the process 4 known in EP 1 538 191 Al is problematical. Furthermore, 5 process stability may be set only with difficulty. 6 Furthermore, the previously described process is 7 distinguished by a lower heating value yield of the input 8 substances. 9 10 DE 10 2005 056 735 B3 discloses a high-performance chamber 11 mixer for catalytic oil dispersions as a reactor for 12 depolymerization and polymerization of hydrocarbonaceous 13 residues to middle distillate. The energy input and 14 conversion rate take place predominantly in the high 15 performance chamber mixer, wherein the pump efficiency of the 16 high-performance chamber mixer is low, that is to say the 17 energy introduced is for the most part converted into mixing 18 energy and friction energy. This process also has a low 19 process stability. 20 21 The object of the present invention is to provide a process 22 for generating middle distillate from hydrocarbonaceous 23 energy sources which is inexpensive, is of low complexity in 24 terms of the process and ensures firstly high process 25 stability and secondly a high yield of the heating value of 26 the energy sources used. 27 28 For achieving the abovementioned object, a process of the 29 type mentioned at the outset provides that at least one 30 hydrocarbonaceous energy source, optionally at least one 31 catalyst and optionally at least one additive, wherein the 32 additive can be a neutralizer, are fed as input material to a 33 reactor containing a process oil mixture, wherein a process 34 oil mixture stream is removed from the reactor and heated to 35 a process temperature between 150*C and 400'C, preferably 36 between 350*C and 380*C, wherein the thus heated process oil 37 mixture stream is fed to a degasser, wherein, in the WO 2009/086908 PCT/EP2008/010990 4 1 degasser, vaporous middle distillate, namely vaporous 2 hydrocarbon compounds in the boiling range of the middle 3 distillate fraction of mineral oil, is separated or removed 4 from the heated process oil mixture stream and wherein a 5 process oil mixture stream relieved of the vaporous middle 6 distillate is recirculated from the degasser to the process 7 oil mixture present in the reactor. 8 9 The invention first provides heating outside the reactor the 10 process oil mixture stream which was removed from the reactor 11 to temperatures of a maximum of 400*C, preferably of a 12 maximum of 350*C to 380 0 C, in such a manner that the 13 formation of reaction coke is reduced. The heating is 14 performed in a gradient-minimized manner. In this context the 15 process according to the invention provides that, during the 16 heating of the process oil mixture temperature peaks, such as 17 occur during heating of the process oil mixture in the 18 process known from DE 100 49 377 C2 on the tube bundles of 19 the evaporator can be excluded in the heat transfer by a 20 suitable process procedure. During the heating of the process 21 oil mixture, the maximum temperature over the entire flow 22 cross section should always be below 400 0 C, preferably below 23 380*C. By means of the lower coke formation thus effected, 24 the high expenditure on cleaning can be reduced and the 25 maintenance intervals prolonged, which contributes to a high 26 economic efficiency of the process according to the 27 invention. Furthermore, in the case of the invention, a 28 process oil mixture circuit in the actual sense is not 29 provided: in the case of the invention the middle distillate 30 vapor released from the heated process oil mixture is removed 31 in the degasser and only process oil mixture which is 32 relieved from the vaporous middle distillate is recirculated 33 to the reactor. By means of a suitable structural design of 34 the degasser, the yield in the process according to the 35 invention may be significantly increased in comparison with 36 the known processes. 37 WO 2009/086908 PCT/EP2008/010990 5 1 Preferably at least some of the heated process oil mixture 2 stream can be applied from the top into the degasser and 3 divided on internals of the degasser into a multiplicity of 4 substreams, wherein the substreams then flow off in a 5 trickling film flow to the reactor. Preferably, in the 6 degasser, an essentially smooth trickling film flow forms 7 with negligible bubble formation, wherein the substreams of 8 the process oil mixture flow off downwards in a stringlike 9 manner. In connection with the invention, it has surprisingly 10 been found that a calm surface of the trickling film flow 11 contributes to a high yield of the heating value of the 12 energy source used, wherein a droplet-like flowing off of the 13 process oil mixture through the degasser is unwanted and 14 preferably be substantially excluded by appropriate 15 structural design of the internals. Some of the heated 16 process oil mixture stream can also be introduced into the 17 degasser tangentially, preferably below the internals, and 18 flows off in the form of a rotary flow on an inner vessel 19 wall of the degasser downward toward the reactor. By means of 20 the division of the heated process oil mixture stream into a 21 first substream applied from the top into the degasser onto 22 the internals and a second substream introduced tangentially 23 into the degasser below the internals, a large surface area 24 of the process oil mixture is generated in the degasser which 25 leads to a high release of vaporous middle distillate in the 26 degasser. 27 28 According to the device, the degasser accordingly comprises a 29 top dividing space and a bottom degassing space, wherein, in 30 the dividing space, flow-guiding and surface-area-increasing 31 internals are provided for dividing a process oil mixture 32 stream and for increasing the surface area of the process oil 33 mixture stream, wherein, preferably, the process oil mixture 34 stream can be delivered centrally into the dividing space 35 from the top onto the internals. The degassing space, 36 furthermore, can comprise at least one inlet for a process 37 oil mixture stream such that the process oil mixture stream WO 2009/086908 PCT/EP2008/010990 6 1 can be introduced tangentially into the degasser and flows 2 off downward toward the reactor as a rotary flow in an inner 3 vessel wall of the degassing space. The inlet into the 4 degassing space here is preferably arranged below the flow 5 guiding and surface area-enlarging internals of the dividing 6 space. The structure of the degasser according to the 7 invention is distinguished by a high self-cleaning ability 8 and is low-maintenance, wherein maximization of the surface 9 area of the process oil mixture flowing through the degasser 10 is ensured with a correspondingly high yield of vaporous 11 middle distillate. 12 13 The main energy input during the heating of the process oil 14 mixture stream removed from the reactor to a process 15 temperature of preferably between 350*C and 380*C proceeds 16 according to the invention by indirect heat transfer from a 17 preferably liquid heat carrier in at least one static mixer 18 having an integrated heat transfer appliance. According to 19 the device, the static mixer can be constructed as a mixing 20 heat exchanger having a multiplicity of tube bundles for a 21 heat carrier, in particular a thermal oil, and mixing 22 elements between the tube bundles for turbulent mixing of the 23 process oil mixture. Therefore, a heating and intensive 24 mixing of the process oil mixture stream to be heated occur 25 at the same time, by a turbulent mixing of the process oil 26 mixture can develop in the static mixer. 27 28 Furthermore, an indirect heat transfer of the process oil 29 mixture contained in the reactor can be provided, wherein a 30 heat transfer from a preferably liquid heat carrier such as, 31 for example, a hot thermal oil, to the process oil mixture 32 can proceed via an outer wall of the reactor. The thermal oil 33 which can be used for heating the process oil mixture stream 34 in the static mixture and for heating the process oil mixture 35 contained in the reactor should preferably have a maximum 36 temperature of below 400 0 C, in particular below 380*C, in 37 order to avoid or reduce the formation of reaction coke, WO 2009/086908 PCT/EP2008/010990 7 1 which ultimately simplifies the maintenance. 2 3 According to the device the reactor can comprise a top 4 cylindrical wall section, wherein, preferably, the top wall 5 section is constructed as a double-shell cylinder having a 6 reactor inner wall and a reactor outer wall and wherein, more 7 preferably, in the double shell, a guide appliance which is 8 spirally mounted on at least one reactor wall is provided for 9 a heat carrier. The top wall section comprises a top inlet 10 port and a bottom inlet port for a heat carrier, wherein the 11 heat carrier flows spirally downward along the reactor inner 12 wall. An additional energy input or else cooling of the 13 process oil mixture is thereby possible in the reactor. 14 15 The process oil mixture stream which is vapor-relieved of the 16 middle distillate can be deflected on entry into the reactor 17 from the degasser, wherein, preferably, a tangential rotary 18 flow on the reactor wall is generated. The process oil 19 mixture in the reactor is in this case subjected to static 20 mixing. 21 22 In a top entry region of the reactor, preferably, internals 23 for deflecting the flow of the relieved process oil mixture 24 recirculated from the degasser to the reactor are provided, 25 wherein the internals are constructed for generating a 26 tangential wall flow along the reactor wall. The reactor is 27 thereby constructed as a static mixer, wherein no active 28 stirring appliances are required. This contributes to an 29 inexpensive structure of the reactor. The reactor can have an 30 inwardly domed vessel bottom, such that a sedimentation cone 31 forms in the bottom region of the reactor, which simplifies 32 the discharge of spent catalyst material, additives and 33 unreacted energy source from the reactor. 34 35 In a further embodiment of the process according to the 36 invention, it is possible provide that a further process oil 37 mixture stream is passed from the reactor into a prereactor WO 2009/086908 PCT/EP2008/010990 8 1 having mixing appliances, wherein the input material is fed 2 to the prereactor and mixed with the further process oil 3 mixture stream in the prereactor and wherein the resultant 4 hydrocarbon-rich process oil mixture stream is recirculated 5 from the prereactor to the reactor. In the prereactor, 6 predewatering and predegassing occur and a catalytic reaction 7 occurs only to a small extent. In the prereactor, the input 8 material is mixed with process oil mixture which is 9 approximately 350*C and originates from the (main) reactor, 10 wherein the liquefaction process of the energy source is 11 initiated. The cracking of hydrocarbon compounds, however, is 12 preferably substantially prevented owing to short residence 13 times in the prereactor and then takes place only in the 14 (main) reactor. The prereactor which is preferably 15 constructed as a screw conveyor comprises at least one feed 16 screw, preferably a double screw as feed unit, for the input 17 material and a mixing vessel connected to the feed screw, 18 wherein, more preferably, the feed screw engages as far as to 19 the bottom region of the mixing vessel and comprises mixing 20 flights at the bottom end. This ensures, firstly, an 21 intensive mixing of the input material with the process oil 22 mixture originating from the (main) reactor and secondly 23 ensures good self-cleaning of the feed screw. 24 25 In principle, the feed screw is cooled by following input 26 material, wherein, however, in particular when the process is 27 being run down, for reasons of material endurance, cooling of 28 the feed screw can be necessary. In principle, it is also 29 possible that heating of the feed screw is provided in order 30 to ensure a sufficiently high temperature in the prereactor. 31 32 For intense mixing, the mixing vessel of the prereactor can 33 comprise at least one bottom inlet for the further process 34 oil mixture stream from the (main) reactor and at least one 35 top outlet for the hydrocarbon-rich process oil mixture 36 stream. The mixing vessel is thereby constructed as a static 37 mixer in which, however, essentially no cracking processes of WO 2009/086908 PCT/EP2008/010990 9 1 the energy source take place. Corresponding internals can be 2 provided as a supplement, in order to intensify the mixing. 3 For the same purpose, a tangential feed of the further 4 process oil mixture stream into the mixing vessel can be 5 provided. 6 7 Finally, a carrier oil can also be fed to the (main) reactor 8 via the prereactor, in particular via the mixing vessel, 9 which carrier oil forms a component of the process oil 10 mixture in the reactor. 11 12 In order to ensure high process stability and a high yield of 13 the heating value of the input material, the volume ratio of 14 the process oil mixture in the (main) reactor to the further 15 process oil mixture in the prereactor should be set to 5:1 to 16 8:1. This assumes an appropriate structural design of the 17 reactor vessel and the mixing chamber of the prereactor. 18 19 The hydrocarbon-rich process oil mixture stream recirculated 20 from the prereactor is mixed with the process oil mixture 21 contained in the reactor and the process oil mixture stream 22 from the degasser which is relieved of the vaporous middle 23 distillate. The hydrocarbon-rich process oil mixture stream 24 is fed into the reactor beneath the internals which are 25 provided in the top region of the reactor for deflecting the 26 flow of the relieved process oil mixture which is 27 recirculated from the degasser into the reactor. The 28 hydrocarbon-rich process oil mixture stream recirculated from 29 the prereactor is preferably in this case introduced 30 tangentially into a mixing zone of the reactor in such a 31 manner that a rotary flow of the entire process oil mixture 32 forms in the reactor. By means of the targeted input of the 33 hydrocarbon-rich process oil mixture stream recirculated from 34 the prereactor, the process oil mixture in the reactor is 35 made to rotate. The direction of rotation of the relieved 36 process oil mixture stream recirculated from the degasser, 37 after entry into the reactor, can in this case correspond to WO 2009/086908 PCT/EP2008/010990 10 1 the direction of rotation of the tangentially introduced 2 hydrocarbon-rich process oil mixture stream from the 3 prereactor. 4 5 According to the device, the reactor can comprise a bottom 6 part having a conically tapering top wall section and a 7 conically tapering bottom wall section, wherein the top and 8 the bottom wall sections are connected to one another by a 9 cylindrical wall section. The process oil mixture stream 10 which is fed to the static mixer for heating and mixing can 11 be withdrawn in the top region of the conically tapering top 12 wall section, wherein there, at least one outlet is provided. 13 By means of this structure of the reactor it is possible to 14 remove, from a top first sedimentation zone of the reactor, 15 the process oil mixture stream which is to be heated and to 16 pass it to the static mixer having an integrated heat 17 transfer appliance. 18 19 In the top region of the conically tapering bottom wall 20 section of the bottom part, at least one further outlet can 21 be provided. This outlet is provided for draining off from a 22 bottom second sedimentation zone of the reactor a process oil 23 mixture stream which is enriched with at least one catalyst 24 and optionally with at least one additive. 25 26 For repeated use of the catalyst it is possible to mix the 27 process oil mixture stream from the top first sedimentation 28 zone, which process oil mixture stream is to be heated, with 29 a process oil mixture stream enriched with catalyst and 30 optionally additive from a bottom second sedimentation zone 31 of the reactor and thereby set a defined catalyst 32 concentration in the process oil mixture. The two streams are 33 mixed before entry into the static mixer, so that both 34 streams are intensively mixed and heated in the mixer. 35 Furthermore, an open-loop or closed-loop control appliance 36 can be provided for open-loop or closed-loop control of the 37 volumetric flow ratio of the process oil mixture stream to be WO 2009/086908 PCT/EP2008/010990 11 1 heated to the enriched process oil mixture stream. 2 3 A substream of the process oil mixture stream to be heated 4 and optionally a further substream of the process oil mixture 5 stream enriched with catalyst and optionally neutralizer form 6 the further process oil mixture stream fed to the prereactor. 7 8 Shortly before transport of the process oil mixture stream 9 which is to be heated and optionally of the process oil 10 mixture stream enriched with catalyst into the static mixer, 11 at least one unconsumed catalyst and/or optionally at least 12 one additive can be added from appropriate reservoir 13 containers. In this case the catalyst and/or additive, before 14 addition, is preferably mixed with a carrier oil or 15 emulsified in a carrier oil, which simplifies the mixing. 16 17 According to the invention it is further provided that the 18 energy source, the catalyst, and optionally the additive 19 which together can form the input material for the process, 20 are mixed with one another before feed into the prereactor 21 and are heated to a temperature of below 120 0 C, preferably to 22 approximately 80 to 100'C. In this case, drying and also 23 aggregate formation occur prior to feed of the input material 24 into the prereactor. The energy source in this case is mixed 25 and heated dry with preferably pulverulent catalyst and/or 26 neutralizer, wherein the resultant aggregate has a high 27 reaction surface area and separation does not take place. 28 Furthermore, the aggregate has a longer residence time in the 29 process oil mixture. The yield is further increased thereby. 30 31 The invention permits individual concepts of the invention to 32 be combined with one another, even if this is not described 33 individually. Furthermore, an independent inventive meaning 34 is ascribed to the static mixing and increase in surface area 35 of the process oil mixture in the degasser and in the 36 reactor, and also to the premixing of the input material with 37 the process oil mixture in the prereactor, wherein the WO 2009/086908 PCT/EP2008/010990 12 1 inventive concepts linked hereby can also independently of 2 one another justify an inventive concept. 3 4 The invention will be described hereinafter by way of 5 examples with reference to the drawing. In the drawings 6 7 fig. 1 shows a schematic process flowchart of the feed of a 8 hydrocarbon-rich energy source together with a 9 catalyst and a neutralizer into an oil circuit for 10 generating vaporous middle distillate and 11 12 fig. 2 shows a schematic process flowchart of the reaction 13 circuit for the generation of middle distillate from 14 hydrocarbonaceous energy sources. 15 16 Fig. I shows a process flowchart which shows the feed of a 17 hydrocarbon-containing energy source 1 into an oil circuit 18 for generating middle distillate 2. The energy source 1, in 19 the present case, is dried and comminuted biomass which is 20 stored in a reservoir vessel 3. The energy source 1 falls 21 owing to gravity from the storage vessel 3 into a first 22 conveyor screw 4. By rotating the spindle, the mixture of 23 matter is pushed into the bottom hopper of a tube chain 24 conveyor 5. The tube chain conveyor 5 transports the energy 25 source 1 to a height of approximately 12 m into a top hopper. 26 From there the energy source 1 falls owing to gravity into a 27 feed screw 6. The feed screw 6 transports the energy source 1 28 at a rate of 5 m 3 /h into a first star feeder lock 7 or into a 29 second star feeder lock 8. The star feeder locks 7, 8 serve 30 for periodic metering of starting material to cone mixers 9, 31 10, wherein each star feeder lock 7, 8 is designed having a 32 transport capacity of 5 m 3 /h. The star feeder locks 7, 8 are 33 dynamic barriers, since material can be transported through 34 and simultaneously a slight underpressure, generated by a 35 vacuum plant, is made possible in the cone mixers 9, 10. 36 37 The cone mixers 9, 10 are degassed in order to decrease the WO 2009/086908 PCT/EP2008/010990 13 1 proportion of oxygen and to keep the risk of ignition of the 2 oil vapor produced in the further process low. The cone 3 mixers 9, 10 have a net volume of approximately 2.4 m 3 . The 4 cone mixers 9, 10 are operated alternately periodically. 5 While the first cone mixer 9 is charged with the energy 6 source 1, the second cone mixer 10 can be mixed using the 7 integrated screw. 8 9 In addition to the energy source 1, periodically at least one 10 catalyst la and/or one additive lb, such as, for example a 11 neutralizer, can be added to the cone mixers 9, 10, where the 12 catalyst la and the additive lb can be a pulverulent mixture. 13 The time of mixing, heating, moisture removal and degassing 14 in the cone mixers 9, 10 is approximately half an hour, and 15 the time period of the charging operation is likewise half an 16 hour. Since both cone mixers 9, 10 have a double shell, 17 heating the mixture of matter in the cone mixers 9, 10 to 18 approximately 100*C is possible. The cone mixers 9, 10 are 19 heated using a heating medium, preferably a thermal oil, such 20 that the input material 12 preferably reaches a temperature 21 of approximately 800C in the cone mixers 9, 10. This leads to 22 drying of the input material 12 with agglomerate formation 23 which has an advantageous effect on the yield in the 24 generation of middle distillate 2 from the energy source 1. 25 The temperature elevation is necessary in order to decrease 26 the proportion of water in the mixture, since it vaporizes at 27 this temperature and can be removed via degassing lines of 28 the cone mixers 9, 10. At an excessive water proportion, a 29 water vapor explosion could occur in further processes. In 30 addition, at a high water proportion the effective separation 31 efficiency in the generation of middle distillate from 32 hydrocarbonaceous energy sources 1 would be decreased. 33 34 The two cone mixers 9, 10 make possible continuous charging 35 of a four-zone reactor 11 shown in fig. 2, wherein, via gas 36 tight slides, the cone mixers 9, 10 are periodically emptied. 37 The input material 12 is discharged from the cone mixers 9, WO 2009/086908 PCT/EP2008/010990 14 1 10, which input material is composed of the energy source 1, 2 optionally the catalyst and optionally at least one additive. 3 The input material 12 passes into a connecting screw 13 and 4 then into a compacting screw 14 in which the input material 5 12 is pressed to half the original size. The connecting screw 6 13 and the compacting screw 14 each have a double shell 7 through which a heating medium, preferably thermal oil, is 8 passed at a temperature of approximately 100 to 1200C. This 9 ensures that the temperature of the input material 12 is kept 10 constant at approximately 1000C. 11 12 Furthermore, the compacting screw 14 has suction points in 13 order to remove further water proportions, inter alia 14 adhering water, of the dried input material 12. In addition, 15 the proportion of oxygen is further decreased. 16 17 From the compacting screw 14, the input material 12 arrives 18 at two ATEX discharge wheels 15, 16. The discharge wheels 15, 19 16 transport the input material 12 into the filling hopper 17 20 of a screw feed mixer 18. The screw feed mixer 18 is a 21 prereactor having a mixing appliance and comprises an oval 22 connecting tube 19, a double screw 20 and an approximately 23 800 1 capacity mixing vessel 21. 24 25 The input material 12 is pushed by means of the double screw 26 20 from the charging hopper 17 through the connecting tube 19 27 into the mixing vessel 21 and mixed with an approximately 28 3500C process oil mixture stream 22 which is withdrawn from 29 the reactor 11 and consists of a carrier oil containing 30 previously dissolved energy source 1 which is in part present 31 in cracked form. The screw ends of the double screw 20 have 32 mixing flights 23 which contribute to the mixing of the input 33 material 12 with the process oil mixture stream 22. The 34 mixing function is supported by metered tangential pumping of 35 the process oil mixture stream 22 from the reactor 11 to the 36 mixing vessel 21 using the spiral housing pump 24, more 37 precisely at two feed points 25, 26 of the mixing vessel 21.
WO 2009/086908 PCT/EP2008/010990 15 1 This ensures a double mixing. In addition, the double screw 2 20 acts as a baffle, since it is situated in the region 3 between the center of the mixing vessel 21 and the wall 4 thereof. By means of the double screw 20, additional 5 turbulence of the flow is caused. The use of a double screw 6 20 is distinguished, furthermore, by a high operational 7 safety at comparatively high temperatures in the mixing 8 vessel 21. 9 10 In the mixing vessel 21 the process oil mixture 22 flows with 11 a rotary motion upwards and mixes with the input material 12 12 which is transported in. After a short time, a hydrocarbon 13 rich process oil mixture stream 26 thus obtained is taken off 14 in the top region of the mixing vessel 21 and recirculated to 15 the reactor 11. 16 17 When the input material 12 arrives in the screw feed mixer 18 18, the liquefaction process begins. The cracking process, 19 namely the cracking of the hydrocarbon chains, owing to a 20 very short residence time of the process oil mixture in the 21 screw feed mixer 18, does not start or only to a slight 22 extent, but starts exclusively or predominantly only in the 23 main process in the reactor 11. Should energy source 1 which 24 is still incompletely dissolved be floating on the surface of 25 the reaction mixture in the mixing vessel 21, it is returned 26 or passed back into the mixture via corresponding internals 27 in the mixing vessel 21. If the input material 12 dissolves 28 in the screw feed mixer 18, residual water fractions are 29 released which are passed out of the screw feed mixer 18. The 30 water vapor passes into a demister 27 which contains packings 31 on which oil droplets which are co-transported by the vapor 32 remain adhering and flow off back into the screw feed mixer 33 18. The water vapor is removed via a vacuum plant and the 34 residual water is liquefied in a condenser 28. 35 36 Both spindles 29, 30 of the double screw 20 operate in a 37 self-cleaning manner. The spindles 29, 30 are rotatably WO 2009/086908 PCT/EP2008/010990 16 1 mounted at the bottom end at the cone base of the mixing 2 vessel 21 and at the top end by shaft passages of the 3 charging hopper 17. The connecting tube 19 is likewise fitted 4 with a double shell, since temperatures of up to 350 0 C can 5 prevail in the mixing vessel 21. The temperature in the 6 charging hopper 17 must not exceed 100*C, since the discharge 7 wheels 15, 16 are designed to have ATEX protection only up to 8 100 0 C. Should too much heat flow upwards via the connecting 9 tube 19, it can be removed via the double shell, wherein a 10 corresponding cooling medium is passed through the double 11 shell. 12 13 For the introduction of carrier oil 31, such as, for example, 14 dewatered waste oil, into the oil circuit, a heatable vessel 15 32 is provided as a reservoir vessel. Liquid residues can 16 also be introduced into the oil circuit in this manner as 17 energy source. By means of the pump 33, the carrier oil 31 is 18 introduced into the mixing vessel 21. The vessel 32 is 19 charged via a pump from an oil store. Finally, via the screw 20 feed mixer 18, the carrier oil 31 can be fed to the reactor 21 11, for example, in order to compensate for vaporization 22 losses. In addition, a carrier oil stream 34 can be passed 23 from the vessel 32 into vessels 35, 36 shown in fig. 2 for 24 producing a catalyst/additive emulsion. The vessels 35, 36 25 possess charging hoppers in order to facilitate charging with 26 the catalyst la and the additive lb. 27 28 The structure of the feed system shown in fig. 1 makes 29 possible sufficient drying, mixing and deaeration of the 30 input material 12. There is therefore no risk of a water 31 vapor explosion in the reactor 11. Likewise, the ignition of 32 released oil vapor need not be feared. Finally, by means of a 33 low water proportion in the reactor 11, a high separation 34 efficiency is ensured. 35 36 In fig. 2, the main circulation system in the generation of 37 middle distillate 2 from the hydrocarbonaceous energy source WO 2009/086908 PCT/EP2008/010990 17 1 1 is shown. The components of the main circuit or reaction 2 system are the four-zone reactor 11, a degasser 37 and also 3 three mixing heat exchanger pairs 38, 39, 40 and also a 4 multiplicity of pumps and the associated piping. 5 6 In the generation of middle distillate 2 from the 7 hydrocarbonaceous energy source 1, a molecular dehydration, a 8 molecular polymerization and a molecular shortening 9 (depolymerization/conversion to oil) take place at a lower 10 temperature without pressurization, compared with pyrolysis. 11 The process procedure is carried out in the main stream at 12 temperatures between 300 and 400*C and at a slight 13 underpressure of -30 to -100 mbar compared with the ambient 14 pressure. The described process is characterized by a higher 15 yield of the heating value of the energy source 1. If the 16 energy source 1 used is polymer waste, more than 70 to 80% of 17 the hydrocarbons present can be obtained. Furthermore, 18 detoxification of environmentally relevant halogens by 19 binding them in the liquid state as immobilizable salts takes 20 place. 21 22 The hydrocarbon-enriched process oil mixture stream 26 23 originating from the screw feed mixer 18 as prereactor is 24 introduced into the reactor 11. The process oil mixture 54 25 contained in the reactor 11 and having the dissolved energy 26 source 1, optionally the catalyst la, optionally the additive 27 lb and carrier oil, is circulated, wherein, per passage, a 28 resulting amount of vaporous middle distillate formed is 29 transferred into a workup system 41 provided above the 30 degasser 37. The workup system 41 is shown only schematically 31 in fig. 2. The main components of the workup system 41 are a 32 vapor expansion predistillation unit or prerectification 33 unit, a rectification column and also condensers and water 34 separators. In the workup system 41, the vaporous middle 35 distillate is separated by distillation into four groups, 36 namely low boilers (hydrocarbon in the boiling range of 37 kerosene and benzene), middle product (gas oil, namely WO 2009/086908 PCT/EP2008/010990 18 1 hydrocarbon mixture in the boiling range of diesel), high 2 boilers (process oil or carrier oil) and bottom product 3 (distillation residues). 4 5 The reactor 11 is structurally equipped with a double cone 6 shape in the bottom region. The reactor 11 has a top 7 cylindrical wall section 43 having a bottom part 44, wherein 8 the bottom part 44 has a conically tapering top wall section 9 45, a conically tapering bottom wall section 46 and a 10 cylindrical central wall section 47. In the region of the 11 bottom part 44, outlet ports 50, 51 are welded in, and also 12 an outlet port 52 for filter bed material 42, which is a 13 component of the bottom circuit. 14 15 A double shell in the region of the top cylindrical wall 16 section 43 serves for additional heat transfer/cooling with a 17 liquid heat carrier, namely thermal oil. The double shell is 18 fabricated in such a manner that the thermal oil introduced 19 via the top inlet port 48 flows around the reactor 11 via a 20 guide appliance mounted spirally on the reactor outer wall 21 and leaves the double shell at an outlet port 49. 22 Furthermore, the reactor 11 has flow-deflecting internals in 23 the region of the lid thereof. 24 25 The reactor 11 may be divided into four zones I-IV. The top 26 zone I is a gas/vapor zone. Here a small amount of middle 27 distillate vapor flows from the mixing zone II beneath into 28 the degasser 37. In the top zone I, the internals for flow 29 deflection are also arranged. 30 31 In the top section of the mixing zone II, in the region of an 32 inlet port 53, the hydrocarbon-rich process oil mixture 33 stream 26 is introduced tangentially into the mixing zone II 34 and mixes with the process oil mixture 54 which is present 35 there. In addition, in the mixing zone II mixing takes place 36 with a process oil mixture stream 55 from the degasser 37, 37 which process oil mixture stream 55 is relieved from vaporous WO 2009/086908 PCT/EP2008/010990 19 1 middle distillate 2. On account of the tangential 2 introduction procedure, the entire liquid rotates in the 3 reactor 11. The rotary motion is additionally kept in motion 4 via the deflected liquid medium from the degasser 37. The 5 direction of motion of the rotating process oil mixture 54 6 and the relieved process oil mixture stream 55 correspond to 7 one another. 8 9 A sedimentation zone III is a third section of the reactor 11 10 and is situated in the top cone segment. Here some of the 11 process oil mixture 54 is transported using the pumps 57, 58, 12 59 as process oil mixture stream 56 which is to be heated 13 through the ports 50 from the reactor 11 to the three mixing 14 heat exchangers 38, 39, 40. 15 16 In a bottom zone IV of the reactor 11, i.e. in the bottom 17 cone segment, the process oil mixture 54 is enriched with 18 catalyst and additive in a high-boiling hydrocarbon matrix. 19 The bottom zone IV is a second sedimentation zone. From the 20 bottom zone IV, a process oil mixture stream 60 which is 21 enriched with catalyst la and the additive lb is mixed by 22 means of the pumps 61, 62, 63 as needed with the process oil 23 mixture stream 56 which is to be heated. A substream 56a of 24 the process oil mixture stream 56 which is to be heated and a 25 substream 60a of the enriched process oil mixture stream 60 26 form the process oil mixture stream 22 which is passed to the 27 screw feed mixer 18 and in which the energy source 1 is 28 dissolved before its entry into the reactor 11. 29 30 The mixing heat exchangers 38, 39, 40 each consist of two 31 mixing heat exchanger units flanged together, wherein mixing 32 heat exchanger units having the trade name "CSE-XR" from 33 Fluitec are used. Between the tube bundles of the mixing heat 34 exchanger units, mixing elements are welded on which lead to 35 a turbulent mixing of the process oil mixture. 36 37 Shortly before the process oil mixture stream 56 and WO 2009/086908 PCT/EP2008/010990 20 1 optionally the enriched process oil mixture stream 60 are 2 transported into the three mixing heat exchangers 38, 39, 40, 3 optionally catalyst la, and optionally additive lb are added. 4 In the mixing heat exchangers 38, 39, 40, the components are 5 then turbulently mixed and heated to approximately 380 0 C. The 6 heating proceeds via a liquid heat carrier, namely thermal 7 oil, which is fed via inlet port 64 and removed via outlet 8 port 65. Furthermore, the mixing heat exchangers 38, 39, 40 9 have inlet and outlet ports for a cleaning oil and also ports 10 for introducing nitrogen. 11 12 The process oil mixture stream 56 and optionally the enriched 13 process oil mixture stream 60 and also optionally the added 14 catalyst la and optionally the added additive lb arrive as 15 heated process oil mixture stream 67 from the mixing heat 16 exchangers 38, 39, 40 into the top region 66 of the degasser 17 37. In the top region 66, the degasser 37 has a dividing 18 space having flow-guiding and surface-area-enlarging 19 internals for dividing and surface area enlargement of the 20 heated process oil mixture stream 67. The process oil mixture 21 stream 67 is in part preferably applied centrally into the 22 dividing space from the top onto the internals in the top 23 region 66 of the degasser 37. 24 25 Furthermore, the degasser 37 has at least one inlet port for 26 a substream 68 of the heated process oil mixture stream 67, 27 wherein the substream 68 is transported tangentially into the 28 degasser 37 in the top region 66 of the degasser below the 29 internals and flows downwards rotating on the vessel inner 30 wall of the degasser 37. 31 32 From the internals of the degasser 37, the process oil 33 mixture 67 flows down as a trickling film, wherein owing to 34 the fine division, a large surface area is created which 35 facilitates the exit of cracked hydrocarbon chains from the 36 process oil mixture 67. These convert into the vapor phase 37 and flow as vaporous middle distillate 2 to the workup system WO 2009/086908 PCT/EP2008/010990 21 1 41. The thin streams flow off together with the substream 68 2 flowing downwards rotating on the vessel inner wall and pass 3 into the four-phase reactor 11. Shortly after entry, they 4 meet the internals in the gas-vapor zone I of the reactor 11, 5 are deflected, and the circuit begins again from the start. 6 7 The process described with reference to figures 1 and 2 for 8 generating vaporous middle distillate 2 is distinguished by a 9 statically forced absolutely turbulent mixing of the process 10 oil mixture in the mixing heat exchangers 38, 39 and 40. This 11 minimizes the heat transfer gradients and the system is fed 12 in a self-cleaning manner with the process oil solids mixture 13 (catalysts, mineral additives). 14 15 As additives lb, lime hydrate, soda, clay flours and 16 bentonites can be used. As catalyst la, preferably mineral 17 zeolite solids are used. The preferably continuous solids 18 addition of catalyst la and/or additive lb is in the range 19 from 0.5 to 20% by weight with respect to the process oil 20 mixture 54 in the reactor 11. Catalysts la and additives lb 21 such as soda and lime hydrate are generally transported at a 22 portion of 1 to 10% by weight, preferably 1 to 5% by weight, 23 to the material feed of the energy source 1 into the reactor 24 1. 25 26 In the sedimentation zone III of the reactor 11, partially 27 undissolved waste material, the catalyst la and additives lb 28 sediment. The catalyst mixed bed thus formed in the bottom 29 zone IV is by means of at least one volumetrically operating 30 pump 69 by transport back via the top part of the bottom zone 31 IV to the mixed catalyst fluidized bed. Thereby, optionally 32 the catalyst la and the additive lb can be repeatedly 33 utilized for the material reaction. The mixed fluidized bed 34 is kept at a constant height by partial discharge of filter 35 bed material 42 by means of the pump 69. A drain-off tank 70 36 is provided for running down the process. 37